This invention relates in general to liquefied petroleum gas recovery and, in particular to improved recovery of liquefied petroleum gas from a raw natural gas feed stream in a cryogenic turbo expander plant.
Propane markets have driven strong demands in the industry for increasing efficiency in the recovery of liquefied petroleum gas. Efficiency in the recovery of liquefied petroleum gas from a raw natural gas feed stream can be measured by the propane recovery yield relative to the capital cost and energy consumption in the recovery process.
To recover propane and heavier hydrocarbon components from a raw natural gas stream, the propane and heavier hydrocarbon components are absorbed and/or liquefied and separated from the more volatile methane, ethane and inert components of the raw natural gas stream. A cryogenic turbo expander plant expends the potential energy of the pressurized inlet raw natural gas, and in some cases, external energy in the form of mechanical refrigeration, to cool and partly condense the raw inlet gas stream. Indirect heat exchange, primarily upstream of the turbo expander, may be used to assist in cooling the inlet raw natural gas stream. In addition, mechanical refrigeration may also be used to assist in the cooling of the inlet gas. As the inlet gas stream cools the heavier, less volatile hydrocarbon components condense first. A two phase separator is provided to separate the condensed liquid phase from the gaseous phase. The remaining more volatile components still in the vapor phase, are fed to the turbo expander. At the turbo expander, the potential energy of the pressurized gas stream is expended to produce mechanical work. This mechanical work is typically utilized to compress residue gas prior to the residue gas exiting the cryogenic plant, or, alternatively, to compress the inlet raw natural gas stream, increasing the potential energy of the inlet raw natural gas. The pressure and enthalpy of the gas is reduced across the turbo expander turbine, thus causing the gas to further cool (to cryogenic temperatures) and condense. As a result, the more volatile components, including a portion of the methane and ethane components condense. Typically, at this stage, greater than 90% of the propane contained in the inlet stream has condensed. Down stream of the turbo expander, a fractionation distillation column is applied in an attempt to strip the more volatile components from the liquid phase to produce a propane and heavier hydrocarbon liquid product stream. In addition, the same fractionation distillation column can be adapted to absorb and/or rectify the propane and heavier components from the gaseous phase, in order to produce an overhead gaseous predominately methane and ethane, product stream. To achieve propane recovery levels typically in excess of 90% recovery yield, a second cold reflux distillation absorber column is applied.
Although liquefied petroleum gas recovery processes capable of high propane recovery levels have been disclosed, the rate of return for the recovery yield has not been economical. Therefore, industry demands for ultra high recovery have not been met with an economical solution. The competitiveness of the petroleum industry has steadily brought about recent design evolutions, thus increasing plant design targets for propane recovery yields. Typically, recent plant designs have targeted approximately 95% propane recovery.
Exemplary cryogenic expander plants and processes are disclosed in Canadian Patent Nos. 1,288,682 (U.S. Pat. No. RE33408), 1,249,769 (U.S. Pat. No. 4,617,039) and 2,223,042 (U.S. Pat. No. 5,771,712) and U.S. Pat. Nos. 5,799,507, and 6,311,516.
Canadian Patent No. 1,288,682 to Khan et al. teaches the utilization of a second cold reflux distillation absorber column, referred as a direct heat exchanger, to absorb additional propane from residual vapor phase on the discharge of the turbo expander. Khan et al. teach that increased percentages of propane and heavier hydrocarbon components can be recovered by contacting the vapor from a gaseous feed stream with at least a portion of the liquefied overhead from the deethanizer.
U.S. Pat. No. 4,617,039 to Loren L. Buck teaches a similar process to recover additional propane from the expander outlet vapor. Buck teaches that the overhead vapor from the deethanizer column is partly condensed and then the liquid condensate is combined with the vapor from the partially condensed feed gases in the deethanizer feed separator which acts as an absorber.
U.S. Pat. Nos. 5,771,712, 5,799,507, and 5,799,507, and 6,311,516. U.S. Pat. No. 6,311,516 disclose other process arrangements applying a similar second cold reflux distillation absorber column.
These processes suffer from characteristics that physically or economically limit propane recovery capability. The increased energy input required to achieve higher levels of propane recovery makes these processes uneconomical. Many of these processes are inherently expensive on a capital cost basis while others require a larger capital expenditure in the attempt to achieve ultra high propane recovery yield. For example, in many processes, expensive stainless steel construction of piping and equipment is required, instead of carbon steel, for cryogenic operation. Still other processes are highly complex and require multiple indirect heat exchangers. These characteristics negatively affect overall recovery efficiency in attempting to achieve ultra high propane recovery yield.
It is an object of an aspect of the present invention to provide an improved cryogenic turbo expander plant process for recovery of liquefied petroleum gas (LPG) (ie. propane and heavier hydrocarbons), as a liquid product, from a raw natural gas feed stream. In a particular aspect of the present invention, the improved cryogenic turbo expander plant realizes an improved efficiency of LPG recovery in relation to associated capital cost and energy consumption.
In an aspect of the present invention, there is provided a process for recovery of liquefied petroleum gas from a feed stream. The process includes:
passing the feed stream through an indirect heat exchanger;
separating the feed stream into a first vapor fraction and a first liquid fraction;
transferring the first liquid fraction to the indirect heat exchanger;
transferring the first vapor fraction to a direct heat exchanger absorber column;
transferring the first liquid fraction to a distilling unit;
distilling the first liquid fraction in the distilling unit to yield a second vapor fraction and a second liquid fraction;
cooling the second vapor fraction in the indirect heat exchanger;
separating the second vapor fraction into a third vapor fraction and a third liquid fraction;
returning at least a portion of the third liquid fraction to the distilling unit;
passing the third vapor fraction through the indirect heat exchanger, at least a portion of the third vapor fraction condensing to a liquid phase;
decreasing pressure of the third vapor fraction such that at least a portion of the liquid phase flashes;
transferring the third vapor fraction to the direct heat exchanger absorber column such that the third vapor fraction mixes with the first vapor fraction, yielding a fourth vapor fraction and a fourth liquid fraction;
transferring the fourth liquid fraction to the indirect heat exchanger;
transferring the fourth liquid fraction to the distilling unit to distill the fourth liquid fraction; and
transferring thee fourth vapor fraction to the indirect heat exchanger, such that, the feed stream exchanges heat with the first liquid fraction, the fourth vapor fraction, and the fourth liquid fraction, all four streams being in parallel. Also, the third vapor fraction exchanges heat with the fourth vapor fraction and the fourth liquid fraction, all three streams being in parallel and the second vapor fraction exchanges heat with, the fourth vapor fraction and the fourth liquid fraction, all three streams being in parallel. Heat is also exchanged between the feed stream and the fourth liquid fraction, after the fourth liquid fraction has exchanged first with the third vapor fraction, and then with the second vapor fraction.
In one aspect, the present invention provides a process with a calculated propane recovery level of about 99.96% with a marginal increase in capital cost, and a decrease in energy consumption compared to prior art processes. Advantageously, recovery of the same level of LPG is possible with lower capital cost or lower energy consumption or both , in comparison to the prior art processes. The economic balance between a lower capital cost plant, lower energy consumption, or higher LPG recovery is different for each particular application.
In another aspect of the present invention, the first and second section of the indirect heat exchanger are incorporated into one plate-fin exchanger up to a plant capacity of about 7.0xc3x97106 std m3/d. Advantageously, this reduces the number of exchangers and reduces interconnecting piping, supports, foundations, and plot spacing. This also reduces the number of cold boxes used for insulating exchangers and interconnecting piping.
In another aspect, heat is exchanged in parallel in all of the streams, rather than in series or in only some of the streams. This provides the ability to exchange additional heat (energy) in the indirect heat exchangers, since temperature approach pinches between the cooling and heating streams are inhibited by applying the parallel heat exchange method within the indirect heat exchanger which distributes the heat transfer with a more linear temperature profile. In turn, recovery levels are increased relative to energy input, thus improving process efficiency. Alternatively, energy input is decreased for a targeted recovery level.
Advantageously, there is less overall capital cost for the construction of the plant since less expensive carbon steel can be utilized, in lieu of stainless steel for the deethanizer column, and the overhead condenser system (ie. deethanizer overhead separator, deethanizer overhead pumps, piping, etc).